Refinery utilizing hydrogen produced from a portion of the feed



United States Patent U.S. Cl. 208-89 5 Claims ABSTRACT OF THE DISCLOSUREA high yield all-hydrocracking overall refinery process for convertingsulfur-containing crude oil to substantially sulfur-free distillateproducts wherein, for example, partial oxidation of a strippedvisbreaker tar is used to provide the hydrogen requirements for atwo-stage hydrocracker treating all of the crude distillates andessentially the only products are sulfur-free distillates and lightproducts with essentially no production of sulfur-containing residualssuch as coke or fuel oil.

CROSS REFERENCES This application is a continuation-in-part of myapplication Ser. No. 601,486, filed Dec. 13, 1966, and now abandoned.

BACKGROUND OF THE INVENTION Field of the invention This inventionrelates to the refining of crude petroleum and like hydrocarbonaceousmaterials of ancient origin. More particularly, the invention relates tohydrocarbon oil refining processes wherein at least a portion of thehydrocarbon oil is upgraded by treating with hydrogen.

Discussion of the present invention in relation to the prior art Thehistorical development of the petroleum refining industry has beencharacterized by a steadily increasing demand for lower boiling, higherquality, distillate prodnets, and decreasing demand for the higherboiling heavy fuel oil components of the crude. In recent years this hasbeen aggravated by restrictions being placed on the burning ofsulfur-containing fuel oils which produce noxious atmospheric-pollutinggases. Consequently, in some areas heavy fuel oil is no longer a salableproduct; and the desirability of converting the crude oil entirely tosulfur-free products becomes apparent.

For reducing the yield from crude of high boiling sulfurcontaining heavyfuel oil, various combinations of treating processes have been devised,including atmospheric and vacuum distillation of the crude, catalyticcracking of the highest boiling distillate portions, thermal crackingand coking of high boiling distillate and residual portions, andhydrofining, hydrocracking, and reforming of the lighter distillatefractions produced by these various treatments. The amount of leftoverresidue can be reduced to a small amount by combinations of thesetechniques, particularly if coking is applied to the highest boilingresidual fraction. In that case, however, there is produced asulfurcontaining coke product which cannot always be disposed of andwhich is similarly subject to the restrictions against burningsulfur-containing fuels.

In accordance with the present invention, it is proposed to separate thecrude oil into two-principal fractions:

3,537,977 Patented Nov. 3, 1970 cracking to produce upgradeddistillates. It is proposed that the first fraction comprise the highestboiling portion of the crude oil and the second fraction comprise thelower boiling portion, which is more amenable to hydrodesulfurizationand hydrocracking than the crude residue. Thus the inventioncontemplates upgrading the distillate portions of crude byhydrogen-consuming reactions and using the nondistillable residue as theraw material for manufacturing the hydrogen to be consumed. The cruderesidue can be converted essentially quantitatively to carbon dioxideand hydrogen in a known manner, by reacting the residual fraction withmolecular oxygen at partial oxidation conditions for synthesis gasgeneration in a hydrogen-manufacturing zone, shift converting carbonmonoxide in the combusted gas mixture of carbon monoxide and hydrogenthereby produced to form carbon dioxide and additional hydrogen, andremoving the carbon dioxide, along with hydrogen sulfide produced fromthe sulfur in the crude residue, to provide high purity hydrogen.

With heavy sulfur-containing crude oils, the usual situation heretoforehas been that the high boiling portion or residue, remaining even aftervacuum distillation of the crude, is more than is needed to manufacturesufficient hydrogen for upgrading of the distillate portions, becausethe refinery schemes have included coking or catalytic or thermalcracking of a portion of the highest boiling distillates and/or therecovery of substantial amounts of straight-run distillates and/ orupgrading of fractions by solvent treatments or like processes notrequiring the use of hydrogen. Hydrogen used in hydrofining the gasolineproduced could be adequately supplied by catalytic reforming. Additionalhydrogen was needed only to desulfurize the cycle oils and straight-runfurnace and diesel oils.

In contrast, the present invention contemplates upgrading essentiallythe entire distillate portion of crude oil entirely by hydrogenation andhydrocracking, to the substantial exclusion of alternate upgradingtechniques which might otherwise have been selected, thus providing anallhydrocracking refinery.

It has been found that the all-hydrocracking concept of the presentinvention gives surprisingly economic superiority over traditionaloverall refinery processing arrangements. Contrary to previousexpectations, upon analysis it was found that the all-hydrocrackingoverall refinery process in accordance with the present invention isnotably superior to, for example, perhaps the currently most widely usedoverall refinery process arrangement which includes a catalytic crackingunit, the most preferable generally being a Fluid Catalytic Crackingunit (FCC). The all-hydrocracking refinery process of the presentinvention is particularly advantageous when there is low demand for fueloil and high demand for jet fuel and motor gasoline. In the comparisonsummarized by Tables I through III below, a generally favorable FCC-hydrocracking process arrangement is compared to the all-hydrocrackingprocess arrangement. For fairness of comparison, the two refineries weredesigned to produce products which are very nearly identical; Table Ishows the calculated products from each refinery as well as the feed.Table II gives a breakdown of the capital investment required for eachrefinery. The grand total at the bottom of Table II shows theall-hydrocracking refinery costs less to build than theFCC-hydrocracking refinery producing the same amount of importatnt majorproducts. Lastly, Table III summarizes the overall economic comparison.From Table II it is seen that the all-hydrocracking refinery requiresless investment than the FCC- hydrocracking refinery, but yet theall-hydrocracking refinery has an income advantage of $6,000,000/year inreduced raw material and operating costs. In addition, the

all-hydrocracki-ng refinery has the capability (flexibility) ofproducing 10,00 b.p.c.d. more jet fuel than the FCC- hydrocrackingrefinery as the latter is more restricted to the production ofautomobile gasoline. The flexibility of the all-hydrocracking refineryto produce more jet fuel is important in view of the rapidly increasingjet fuel market.

TABLE I.STOCK BALANCE COMPARISON 4 THE INVENTION The present inventionprovides a process for converting sulfur-containing crude oilsubstantially entirely to sulfurfree distillate products and gaseousby-products. In accordance with the invention, the crude oil is firstseparated into essentially three fractions: a normally gaseous fraction,a sulfur-containing distillate fraction, and a sulfur-containingresidual fraction. The residual fraction TABLE III.-OVERALL ECONOMLCCOMPARISON Net income advantage of All Hydrocracking vs.FCC-Hydrocracking at 95% yearly operating factor .95365X$11,000/day=$6,000,000/

.Capital investment savings for All Hydrocracking vs. FCC-Hydroeracking.See Table I, 223.7=2l4.0=$9,700,000,

i fi g i q; is reacted with molecular oxygen at partial oxidation conficm b 1O ditions for synthesis gas generation in a hydrogen manufacturingzone, the carbon monoxide produced is shift eb camomii, cmde 154,910147,774, converted to carbon dioxide, and the carbon dioxide and P g g fand 111180 17,964 17,964 hydrogen sulfide are removed to provide highpurity hy- 3 5 equiment fuel O11 6,449 4,549 15 drogen. The entiresulfur-containing distillate fraction is 4,413 332 reacted with aportion of the said high purity hydrogen 3,161 0 in contact with asulfactive hydrogenation catalyst at 2 2 elevated temperature andpressure in a hydroconversion 301930 301930 zone, whereby sulfurcompounds in the distillate are con- 5' i 8 g gs verted to H 8 and thereare formed gaseous by-products '200 '200 and hydrofined distillate. Atleast a portion of the hydro- $38 %28 fined distillate is reacted withthe remaining portion of 3: 100 0 the high purity hydrogen in contactwith an active hydrocracking catalyst at elevated temperature andpressure in another hydroconversion zone, thereby converting hy- 2355 34 51mm drofined distillate to hydrocracked lower boiling distillateTABLE IL-CAPITAL INVESTMENT COMPARISON FCC and hydroeracking Allhydrocracking Capacity Dollars Capaeity- Dollars B/CD (million) B/CD(million) On Plot:

Crude unit 16. 9 15. 8 1.0 1.0 15.3 13.1 12.1 31.2 40.0 5.3 5.3

' Process plts. total 144. 0 136. 5

Utilities:

Steam, lb./hr 650, 000 5. 5 500, 000 4. 0 Water treating, lb./l1r1,000,000 3.0 1,000,000 3.0 Water supply 6 6 Drains and waste disposal2. 0 2. 0 g.p.m 167,000 4.4 178,000 4.0 Relief, fuel gas, fuel oil 1.0 1. 0 Elect. power, kw 101, 000 5. 3 113, 000 5. 9 Tanks and Lines,Etc.:

Tanks and lines (i) 31. 2 (2) 30.2 e :5 1:11:11: I5 Loading racks, bbl.filling and solvent storage 2, 5 Other:

Marine 5. 0 Bldgs 5.0 Site development 5. 5 Land 10.0

Nonprocess total 75. 0 Overhead 2. 2. 5

Grand total 223. 7 214. 0

9.335001. 8.8171 bbl.

and forming gaseous by-products. Hydrogen sulfide is removed from thetreated distillate and by-products to recover gaseous by-products and torecover sulfur-free distillate products comprising the hydrocrackeddistillate and any hydrofined distillate which was not hydrocracked. Theamount of hydrofined distillate which is hydrocracked, and the extent ofconversion, are such that all the hydrogen produced from the residualfraction is consumed in hydroconversion.

The present invention thus provides for balancing the hydrogen consumedin the hydrofining and hydrocracking processes with the hydrogenproduced by the partial oxidation of the residue. In the case of heavycrude oils, it

is often desirable to reduce the quantity of heavy residue remaining toless than a vacuum residuum, because the hydrogen required forsatisfactory upgrading of the distillate portion may be less than can beproduced from a vacuum residue. In one preferred embodiment of theinvention, the crude oil is separated into the three fractions by acombination of distilling the crude oil under vacuum, visbreaking thevacuum distillation bottoms, and stripping gaseous by-products andadditional sulfur-containing distillate from the visbreaker efiluent ina pitch stripper, forming the sulfur-containing residual fraction as thepitch stripper bottoms which feeds the hydrogen plant. In a secondembodiment, the vacuum residuum is processed through a solventdeasphalting unit, and the bottoms from the solvent deasphalter used tomake hydrogen. In a third embodiment, the vacuum residuum is processedthrough resid hydrocracking and the bottoms from this unit used to makehydrogen. In such case, fuel oil production is low or zero; and thenecessary hydrogen is produced from the lowest value hydrocarbon.

While, as noted, in the practice of the invention the use ofsupplementary distillate conversion and upgrading processes notrequiring hydrogen is minimized, it is generally desirable to upgradethe octane number of gasoline boiling range distillate products of theprocess by catalytic reforming of a portion of the sulfur-freedistillate produced. In such case, hydrogen produced in the catalyticreforming can be used to augment the supply of high-purity hydrogenproduced by the partial oxidation of the residual fraction.

Where, as contemplated in the invention, the distillate upgradingprocesses involve essentially only hydrogenconsuming hydrofining andhydrocracking reactions, there can be provided an integrated refineryscheme wherein the net input of energy is minimized. The hydrofining andhydrocracking reactions are strongly exothermic, and thus provide asource of high temperature heat energy. This heat energy can betranslated into work energy and used to supply a major portion of thework energy required in the hydrogen manufacturing zone. In particular,a major cost factor in the production of hydrogen by partial oxidationof residua is the large amount of energy required to remove the carbondioxide from the synthesis gas and the energy required for providing thehydrogen at the elevated pressures used in the hydrofining andhydrocracking reactions. The partial oxidation process centers the mainpower load of the refinery in the air compressors for the oxygen plant,oxygen compression, and hydrogen recycle. All of these loads can beserved by centrifugal compressors amenable to steam-turbine drives, andbyproduct steam generated in removing the exothermic reaction heat ofthe hydrofining and hydrocracking reactions can be used in automaticextraction condensing steam turbines.

Moreover, the partial oxidation, shift conversion, and hydrogenpurification can be carried out at substantially higher pressures thancan be used in other hydrogen manufacturing processes such assteam-natural gas reforming. The adverse effect of pressure on theequilibrium reaction CH +H O CO'+3H does no enter into consideration atthe conditions for partial oxidation of heavy oils, even though steammay be present, because the extent of conversion is primarily determinedby reaction kinetics. Thus, partial oxidation is being carried out at600 p.s.i.; operations at 1200 psi. have been reported; and it isapparent that even higher pressures of 2000 p.s.i. or more can be usedcorresponding to the pressure levels used in the hydroconversionreactions. At such pressures the unit costs of gas compression andhydrogen purification are greatly reduced, and energy interchange withthe hydroconversion portion of the refinery is facilitated.

It is accordingly part of the concept of the present invention in itspreferred embodiments, that the partial oxidation hydrogen manufacturingis to be carried out at a high pressure above 600 p.s.i.g..Specifically, the pressure should be at least 1200 p.s.i.g., and moredesirably in the neighborhood of 15002000 p.s.i.g., or higher. Theinvestment cost for a partial oxidation hydrogen plant is substantiallygreater than for a natural gas-steam reforming hydrogen plant, in largepart because there must be provided an air separation unit to supplypure oxygen, and a much larger amount of CO is produced which must beremoved from the H The removal of this CO is a costly operation by theusual techniques involving aqueous amine absorbents employed in lowpressure, below 600 p.s.i.g., hydrogen plants. At the higher pressures,preferably employed in present invention, however, more economical COremoval techniques can be used to greater advantage, involving the useof low temperatures below -10 F. A major part (60-90%) of the CO can becondensed, with the liquid CO when expanded at lower pressures,providing much of the refrigeration. Final cleanup with low viscosity,low freezing point, organic oxygen compounds such as acetone, methanol,and the like as absorbents is greatly facilitated, as the solubility ofCO in such solvents at high pressure and low temperature is greatlyimproved so less solvent must be heated, cooled, and circulated. Powercan be generated from the expanding C0 The cost of gas compression isgreatly reduced since only a relatively small volume of oxygen has to becompressed to provide a much larger volume of hydrogen already at thehigh pressure used. Thus a specific embodiment of the inventioncomprises the combination of high pressure partial oxidation of theresidual portion of crude and low temperature removal of the COproduced, using an organic oxygen compound liquid absorbent, preferablyas a final cleanup after first condensing out part of the CO with highpressure hydroconversion of the nonresidual portion of crude using thehydrogen so produced.

BRIEF DESCRIPTION OF THE DRAWING Referring now to the attached drawingto further explain the invention, there is shown a simplified block flowdiagram illustrating schematically the steps involved in the simplifiedrefinery of the present invention.

DETAILED DESCRIPTION OF THE DRAWING As shown in the drawing, a heavysulfur-containing crude oil in line 11 is passed to crude distillationfacilities in zone 12, wherein straight run normally gaseous com ponentsof the crude are withdrawn as an overhead fraction in line 13. Adistillate fraction comprising essentially all of the normally liquiddistillate portion of the crude, and having an end boiling point of atleast 1000 F., is

withdrawn in line 14, leaving a vacuum residuum which a is withdrawnthrough line 15. The vacuum residuum passes to visbreaker 16 wherein itis mildly thermally cracked in a known manner producing gaseousby-products and additional distillates with minimal coke formation. Theefiluent of the visbreaking zone passes via line 17 to pitch stripper 18wherein the gaseous and distillate portions are stripped out andreturned via line 19 to the crude distillation facilities. The remainingvisbroken pitch stripper tar in line 20 is withdrawn as feed to apartial oxidation process for hydrogen manufacture in hydrogenmanufacturing zone 21.

In hydrogen manufacturing zone 21 the stripped tar bottoms is reactedwith molecular oxygen, nitrogen-free, supplied from an air fractionationunit through line 22. The reaction is carried out in a noncatalyticcombustion zone at temperatures of 20003000 F. using a stoichiometricdeficiency of oxygen such that the combusted gas mixture comprisespredominantly carbon monoxide and hydrogen, containing a small amount ofelemental carbon, some carbon dioxide, hydrogen sulfide, and only asmall portion of unreacted hydrocarbon. The combusted gas mixture isthen worked up to provide high purity hydrogen in a known manner,involving removal of the carbon particles, shift conversion of thecarbon monoxide by reaction with Water to form carbon dioxide andhydrogen, and removal of the carbon dioxide and, usually, any remainingcarbon monoxide. Sulfur compounds in the stripped tar bottoms appear ashydrogen sulfide, which is removed separately or with the carbondioxide. These by-products may be withdrawn as a mixture, but preferablyat least a major portion of the carbon dioxide is withdrawn as by line23 as a high purity CO stream which can be vented to the atmosphere, andthere is also obtained a mixed gas stream of CO and H 8 in line 24having a sufliciently high H 8 concentration to serve as the feed gas toa sulfur recovery plant. This H S and CO removal and separation can beaccomplished by known processes, using for example a methanol absorbent.

The high purity hydrogen recovered in line 25 is divided intoessentially two positions, a portion in line 26 being reacted with thedistillate fraction of the crude petroleum in hydroconversion zone 28,and the remaining portion of high purity hydrogen in line 27 beingreacted with a portion of hydrofined distillate in hydrocrackingreaction zone 29.

In hydroconversion zone 28 the distillate fraction and the portion ofhigh purity hydrogen, generally together with hydrogen-rich recycle gas,are passed into contact with a sulfactive hydrogenation catalyst atelevated temperatures in the range 650-850 F. and pressures in the rangeof 1500-4000 p.s.i.g., at throughput rates, based on the distillatefraction, of 02-10 volumes of oil per hour per volume of catalyst (LHSV)and, based .on the total fresh and recycle hydrogen, of 200020,00-0standard cubic feet of hydrogen-rich gas per barrel of oil. Net hydrogenconsumption will be in the range 500-2000 s.c.f./bbl. The catalystemployed desirably comprises a Group VI metal or compound thereof, suchas the sulfide, and a Group VIII metal or compound thereof, such as thesulfide, associated with a porous inorganic oxide carrier providing highsurface area, such as alumina or combinations of alumina with silica,magnesia, zirconia, titania and the like. The catalysts are of the typewhich selectively promotes hydrodesulfurization and hydrodenitrogenationreactions, converting sulfur and nitrogen compounds in the oil to H Sand NH to form hydrofined distillate which is sufiiciently sulfurandnitrogenfree for recovery as distillate products, or at leastsutficiently purified for the hereinafter-described subsequenthydrocracking. Typical suitable catalysts comprise nickel or cobalttogether with molybdenum or tungsten dis persed in or on alumina,alumina with up to 50% silica, silica-magnesia,alumina-silica-zirconia-titania, etc. The catalysts can be prepared byknown impregnation and coprecipitation methods.

The hydrofined distillate produced is passed via line 30 to distillationand recovery facilities in zone 31 wherein gaseous by-products areseparated and removed through line 32, and the distillate can, ifdesired, be separated into a gasoline boiling range fraction in line 33,a kerosene or jet fuel boiling range fraction in line 34, a middledistillate gas-oil fraction in line 35, and a higher boiling gas-oilfraction in line 36. Alternately, all or a portion of the hydrofineddistillate may be passed via line 37 to zone 29, but it is usuallydesirable to first remove part of the gaseous by-producsts.

At least a portion of the hydrofined distillate, desirably including thehighest boiling portion, is passed as by line 36 and/or line 37 tohydrocracking reaction zone 29, wherein the hydrofined distillate andthe remaining portion of high purity hydrogen in line 29, together withhydrogen-rich recycle gas, are passed into contact with an activehydrocracking catalyst at tempjeratures in the range 400-800 F.,pressures in the range l000-4000 p.s.i.g., at flow rates of oil relativeto catalyst, flow rates of hydrogen relative to oil, and hydrogenconsumption in the same ranges as described for hydroconversion zone 28.The catalyst employed in hydrocracking zone 29 desirably is a catalystof the type exhibiting substantially greater and more acidichydrocrackling activity than the type used in zone 28. Generally, thecatalyst will comprise a Group VIII metal or compound thereof, such asthe sulfide, e.g., cobalt, nickel, platinum, or palladium, associatedwith a porous inorganic oxide carrier having the properties of an activecracking catalyst, such as silica-alumina comprising 75-90% silica. Thecarrier may be, or may contain, aluminosilicate molecular sieves such ashave been incorporated in cracking catalysts. Methods of preparing thecatalysts are well known. The reactions promoted by this type ofcatalyst are of the type invloving a carbonium ion mechanism, wherein acharacteristic is the production of a butane-rich fraction having ahigher than equilbrium ratio of isobutane to normal butane accompanyingthe hydrocracking of the hydrofined fraction to hydrocracked lowerboiling distillate of improved properties. To take maximum advantage inzone 29 of the activity of the hydrocracking catalysts, thehydroconversion in zone 28 is carried to the point of eliminatingnitrogen compounds in the hydrofined distillate to below about 10 p.p.m.In many cases, however, it is advantageous to permit up to about 200ppm. organic nitrogen to remain in the oil, and removal of the NH;;formed is then not essential.

The hydrocracked distillate eflluent of zone 29 is returned via line 38to the distillation and recovery zone 31 for removal of gaseousby-products in line 32. The sulfur-free distillate roducts recovered inlines 33, 34, and 35 thus comprise the hydrocracked distillate and anyhydrofined distillate from zone 28 which was not hydrocracked in zone29.

The gaseous by-products in line 32 generally contain hydrogen sulfideproduced in the hydrofining reactions occurring in zone 28, and thegaseous fraction of line 13 from crude distillation zone 12 will alsocontain hydrogen sulfide liberated in the visbreaking and pitchstripping operations. These gaesous streams are passed to purificationzone 39 wherein the hydrogen sulfide is removed and withdrawn throughline 40, providing a sulfur-free light gas stream in line 41. The lightgases can be utilized as sulfur-free fuel gas, including, desirably, therecovery of a portion thereof as liquefied petroleum gas. Ammonia and HS formed in zones 28 and 29 may be removed by water scrubbing, and maybe recovered separately from the water in a manner now known.

The above-described simplified refinery scheme wherein the upgrading ofthe distillate portion of crude is carried out substantially entirely byhydrogen-consuming conversion reactions, with the hydrogen beingprovided by a partial oxidation which eliminates the residual portion ofthe crude, can be made self-balancing with respect to hydrogenproduction and consumption. Thus, in the case of a light, high gravity,crude oil feed, the fraction of crude separable as distillate in line 14is greater relative to the amount of residue in line 20 as compared tothe situation where the crude oil feed is a low gravity heavy crude.With the higher gravity crudes there is accordingly less residueavailable for use in manufacturing hydrogen. The distillates from highgravity crudes, however, usually require less hydrogen for upgrading asthey tend inherently to require less severe hydrofining andhydrocracking. Conversely, in the case of low gravity crudes there is alarger amount of crude residue remaining for conversion to hydrogen, andmore hydrogen is needed to upgrade the heavy, and usually highlycontaminated, distillate fraction. Further, it will be recognized thatthe visbreaking and pitch stripping operations provide additionalflexibility for providing more or less distillate via line 19 and moreor less residue via line 20.

The extent of upgrading by hydrofining and hydrocracking reactions in aconventional refinery scheme, wherein these processes were heretoforeapplied primarily as means of upgrading catalytic, coker, or thermalcycle oils, is either limited by the availability of hydrogen producedby catalytic reforming, or limited by the availability and cost ofnatural gas where the hydrogen is produced by steam-methane reforming.Accordingly, carrying the hydroconversion reactions to the optimumconversion for maximum upgrading was rarely practiced. In the presentinvention, however, the partial oxidation hydrogen manufacturing zoneserves also as a means of disposing of the highest boiling residue ofthe crude by burning, wherein the H 5 produced is recovered, thusfurther justifying the use of partial oxidation in preference to otherhydrogen manufacturing processes. In such a case, the availability ofhydrogen becomes large enough such that the use of hydrofining andhydrocracking reactions as the sole means of upgrading the distillateportion of crude requiring desulfurization becomes practical anddesirable. Furthermore, the hydroconversion reactions can be carried outat maximum hydrogen consumption for providing optimum upgrading to formtruly sulfur-free distillate products.

The following example specifically illustrates one method of practicingthe invention, and further describes details of various features in thesteps involved. In the example, a crude oil is converted at 96% yield tosulfurfre e C liquid products boiling below 650 F., with zero productionof heavy fuel oil.

EXAMPLE A heavy crude oil, 14 API, containing 1.1 weight percent sulfurand 0.77 weight percent nitrogen, is desalted, preheated, and fed to avacuum distillation column operated at 2 p.s.i.a. An overhead fractionis withdrawn through the jet ejectors, cooled, and separated in a sealdrum at about 2 p.s.i.g. and 90100 F. into a normally gaseous fractionand condensed oil fraction having an end point of about 550 F. From thevacuum crude column there is withdrawn a top side out boiling from about550 to 650 F., a portion of which is cooled and returned as top reflux,and another minor portion of which is withdrawn as a straight-run dieselfuel. A lower side out boiling from about 650 to 1050 F. is withdrawnfrom the vacuum crude column, and another heavier side out boiling fromabout 850 to 1050 F. is withdrawn from lower in the column. The sidecuts are cooled by heat exchange to the crude oil feed, therebyproviding part of the feed preheat, and the 550650 F. and 6501050 F.side cuts are then combined with the fraction condensed from theoverhead to form a single distillate fraction. The vacuum column bottomsamounts to about 30 volume percent of the crude oil feed, and a smallportion of the bottoms is combined with a portion of the 8501050 F.distillate cut for asphalt blending, the remainder of this distillatereturning to the vacuum column for internal reflux. The remainder of thevacuum column bottoms, boiling substantially entirely above 1050 F., isfed to a visbreaker.

The vacuum column bottoms is heated in the visbreaker furnaces to about900 F. at 200250 p.s.i.g., which results in cracking to gas-oildistillate and lighter products in a yield of about 53 volume percent.The furnace effluent is quenched to about 750 F. to stop the reactionand then flashed sequentially in a high pressure flash drum and in a lowpressure flash drum. Bottoms from the low pressure flash are heated toabout 780 F. and passed to a tar stripper, or pitch stripper, whereindistillates are stripped out with superheated steam at about 3 p.s.i.a.The gases and vaporized distillates from the flash drums and tarstripper are returned to the vacuum column, whereby the normally gaseousby-products appear in the vacuum column overhead, and the normallyliquid distillates appear ultimately in the single distillate fractionformed from the vacuum side cuts. The tar bottoms from the pitchstripper amounts to about 15 percent of the original crude oil, and thistar is fed to the hydrogen manufacturing unit.

In the hydrogen manufacturing unit the tar bottoms are dispersed insteam, preheated, and reacted with pure oxygen, supplied from an airfractionation plant, in a ceramic-lined reaction zone at a pressure of1700 p.s.i.g. and a temperature of about 2600 F., the proportions ofoxygen, steam, and tar bottoms being regulated such that thistemperature is autogenously maintained. The resulting combusted gasmixture contains about 4050% hydrogen, about 4050% carbon monoxide,about 1 to 4% elemental carbon, about 310% carbon dioxide, less thanabout 1% methane, H 8, and N The hot gases leaving the partial oxidationreaction zone are immediately quenched with liquid water, which scrubsout the entrained carbon particles. Additional steam is added to thescrubbed gas mixture of carbon monoxide, hydrogen, hydrogen sulfide andcarbon dioxide, and the mixture is passed into contact with asulfur-tolerant shift conversion catalyst at a temperature of about700-900 F. to convert the carbon monoxide to carbon dioxide with thesimultaneous production of hydrogen. Carbon dioxide, unreacted steam,carbon monoxide and other impurities, including hydrogen sulfide formedfrom the sulfur compounds in the stripped tar bottoms, are removed fromthe shifted gas to produce high purity, -99%, hydrogen. (The nitrogencompounds in the tar appear as an N; diluent in the H The Water isremoved by condensation. The carbon dioxide and hydrogen sulfide areadvantageously removed by partial condensation and by contacting at alow temperature, about -80 F., with liquid methanol. The methanolabsorbent is regenerated in stages to release a pure carbon dioxidestream and a carbon dioxide-hydrogen sulfide mixture having a highenough H 8 concentration (about 12% H S-88%CO to permit its use as feedto a conventional sulfur plant. Unconverted carbon monoxide is removedby catalytic conversion to methane. In the hydrogen manufacturing zonethere is thus produced about 15,000 standard cubic feet of hydrogen perbarrel of stripped tar bottoms fed to the partial oxidation. Thehydrogen is compressed from about 1500 p.s.i.g. to about 3000 p.s.i.g.for use in the hydrofining and hydrocracking reactors.

The distillate fraction comprising the straight-run and visbreakerdistillates is passed to a catalytic hydroconversion zone forhydrodcsulfurization and hydrodenitrogenation, which is accompanied bysome hydrocracking. The blended distillate feed has a gravity of about18.5 API,

boils from 150 to 1100 F., and contains one weight percent sulfur andnearly 5000 p.p.m. nitrogen. The distillate fraction and about 10,000s.c.f. of hydrogen-rich gas per barrel, including makeup hydrogen fromthe hydrogen manufacturing zone, are passed at about 750 F. and 3000p.s.i. downflow through a reactor containing fixed beds of nickelsulfide-tungsten sulfide-alumina-silica catalyst particles at 0.85 LHSV.About 1500 standard cubic feet of hydrogen per barrel of feed isconsumed in the hydroconversion reactions resulting. The eflluent of thereactor is cooled and water washed at the high pressure to remove theammonia and a portion of the hydrogen sulfide formed. The hydrogen-richgas is recycled, and the liquid oil is then flashed at a lower pressureto remove dissolved normally gaseous by-products. The oil is then passedto distillation facilities wherein an overhead cut boiling below 200 F.is separated and liquid distillate cuts are formed, comprising 200 300F. gasoline, 320- 400 F. jet fuel, 400-500 F. kerosene, 500650 F. dieselfuel, and a 650 F.+bottoms fraction. The gasoline cut contains only onep.p.m. sulfur and 0.5 p.p.m. nitrogen. All of the side cuts areessentially sulfur-free, containing less than p.p.m. sulfur, and the 650F.+bottoms cut contains only 3 p.p.m. nitrogen.

The 650 F .+bottoms and a portion of the 500-650 F. diesel fuel out arecombined and passed to a catalytic hydrocracking zone with about 6000standard cubic feet of hydrogen-rich gas per barrel, including makeuphydrogen from the hydrogen manufacturing zone, at about 675 F. and 3000p.s.i.g. to contact in a reactor a nickel sulfidesilica-alumina catalystat 0.5 LHSV. About 60 volume percent of the gross feed is converted todistillates boiling below 500 F. Hydrogen consumption is about 1500s.c.f. per barrel of feed converted. The reactor efiluent is cooled atthe high pressure to separate hydrogen-rich recycle gas, and the oil isthen flashed at a lower pressure to separate dissolved normally gaseoushydrocarbons. The hydrocracked oil is then returned to the distillationfacilities together with the effiuent of the hydroconversion reactor forrecovery of the described products.

The normally gaseous fraction from the vacuum crude column and thegaseous fraction from the distillation of the hydroconversion zone andhydrocracking zone efiiuents are combined and passed to a gas recoveryzone to separate an H s-containing light gas stream from propane,butanes, isopentane, and a C 200 F. cut for gasoline blending. The H 8in the liquid propane and in the light gas stream is absorbed in anaqueous monoethanolamine solution. The recovered H S is stripped fromthe MEA in a reboiled regenerator and is sent to a sulfur plant forconversion to elemental sulfur. Residual H 8 is removed from the propaneproduct to meet LPG specifications by adsorption in a molecular sieveadsorption unit.

From 1000 barrels of heavy crude oil feed there are recovered asproducts about 2 tons of sulfur, 1600 pounds of NH;. 120 barrels ofpropane and butanes, 460 barrels of motor gasoline, about 380 barrels ofjet fuel, and about 120 barrels of stove oil and diesel gas oil. TheZOO-320 F. gasoline cut is passed to a catalytic reformer, and thehydrogen produced therein is combined with the hydrogen produced in thepartial oxidation hydrogen manufacturing unit, reformer hydrogenamounting to about 10% of the total. Heat liberated in thehydroconversion and hydrocracking zones is used to generate steam, andextensive use is made of power recoverey turbines on high pressurestreams, to supply nearly 70% of the power requirements of the partialoxidation hydrogen plant. The remainder of the total power requirementsof the process is provided by a gas turbine generator, the hotcombustion gases from which are exhausted into a high pressure boilersupplying 1700 p.s.i. steam to the partial oxidation unit. Overall steamproduction and consumption are entirely in balance.

For comparison with the foregoing example wherein there is zero heavyfuel oil production and 96% yield of sulfur-free light distillatesbesides propane and butanes, in a prior art refinery scheme applied tothe same crude oil using vacuum crude distillation in combination withcatalytic cracking of gas oils and thermal cracking of residua, theyield of light distillate is only 48% and of heavy fuel oil is 47%. Byadding units for hydrofining and hydrocracking of the cracked cycle gasoils, the light distillate yield is raised to 59% and the fuel oil yieldis lowered to 40%. In a combination of vacuum crude distillation,visbreaking, and pitch stripping with hydrofining and hydrocracking ofthe distillates, wherein the hydrogen is supplied by steam-natural gasreforming, the light distillate yield is 78% and fuel oil 30% (tarbottoms plus gas oil cutter).

It is within the contemplation of the present invention that minorportions of the original crude oil feed or of the gaseous, distillate,or residual fractions may be withdrawn from the process if suitable foruse in products which are not required to have a low sulfur content. Forexample as described in the example, part of the vacuum residuum, or thepitch stripper bottoms, may be used in asphalt manufacture, and part ofa heavy distillate fraction may be used as a blending or cutter stock.Also, a portion of sulfur-containing straight-run distillate fractionmay bypass the hydrofining and hydrocracking zones for blending with asulfur-free distillate product of the process, to provide a fuel oilhaving a satisfactorily low sulfur content. Thus when it is stated thatthe crude oil is converted substantially entirely to sulfur-freedistillate products and gaseous by-products in the process of theinvention, the intended meaning is that no more than 12 about 10% of thecrude or fractions thereof will be other wise processed. In theforegoing example, for instance, the total of streams so withdrawn forasphalt, cutter, and diesel fuel amounts to less than 4% of the crude.

Persons skilled in the art will recognize that it is within the intendedscope of the invention to carry out the crude separation into thedistillate and residual fractions by other means equivalent to themethods specifically described herein. Thus, the crude may be distilledin several steps and multiple columns, including atmospheric and vacuumdistillation columns. Th amount of the residuum fraction may be reducedby other means equivalent to visbreaking and tar or pitch stripping forproviding additional distillate fraction. The hydrofining andhydrocracking may be carried out in a variety of ways besides downfiowof hydrogen and oil through fixed beds of catalyst in reactors. Forexample, there may be used concurrent upflow of gas and oil orcountercurrent flow, and the catalyst may gravitate or be fluidized ormaintained in a liquid slurry. Multiple reactors in series or parallelcan be used, portions of the crude distillate can be separately treated,and separate distillation facilities may be used for the recovery oflight distillates from the hydrofined oil and the hydrocracked oil. Thehydrofined oil may pass directly to the hydrocracking zone. Variousarrangements of multiple columns can be used in working up the products.In the train of steps in the hydrogen manufacturing plant using partialoxidation, any of a variety of known means for removing, recovering, andrecycling the carbon formed can be incorporated. High pressure steam canbe generated, and more than one stage of shift conversion and/or COremoval can be provided. Improved shift conversion catalysts arecommercially availablesuch that a later stage of shift conversion can becarried out at a lower temperature of 400700 F. after an earlier highertemperature stage, and the CO and H S can be removed in stages usingdifferent kown absorbents. Residual CO can be left in the H in somecases, or it can be removed by absorption in cuprous solutions such ascopper ammonium acetate. The specific mode of operating the overallprocess described in detail herein, however, has distinct advantages ofsimplicity, low operating cost, and optimum energy integration uniquelyfulfilling that object of the invention, to provide a simplifiedrefinery.

I claim:

1. A process for converting sulfur-containing crude oil substantiallyentirely to sulfur-free distillate products and gaseous by-products,comprising:

(1) separating the crude oil into essentially three fractions, (a) anormally gaseous fraction, (b) a sulfurcontaining distillate fraction,and (c) a sulfur-containing residual fraction;

(2) reacting the sulfur-containing residual fraction with molecularoxygen at partial oxidation conditions for synthesis gas generation in ahydrogen manufacturing zone to obtain raw hydrogen-rich gas, shiftconverting CO contained in the raw hydrogen-rich gas to CO and thenremoving H S and CO from the raw hydrogen-rich gas to provide highpurity hydro- (3) reacting the sulfur-containing distillate fractionwith a portion of said high purity hydrogen in contact with a sulfactivehydrogenation catalyst at elevated temperature and pressure in ahydroconversion zone, converting sulfur in said distillate to H 5 andforming hydrofined distillate and gaseous by-products;

(4) reacting at least a portion of said hydrofined distillate with theremaining portion of said high purity hydrogen in contact with an activehydrocracking catalyst at elevated temperature and pressure in ahydroconversion zone, converting hydrofined distillate to hydrocrackedlower boiling distillate and forming gaseous by-products;

() removing H 5 from the distillates and the gaseous by-productsobtained from each of the hydroconversion zones, recovering gaseousby-products, and recovering substantially sulfur-free distillateproducts comprising hydrocracked distillate and any hydrofineddistillate not hydrocracked.

2. A process in accordance with claim 1 wherein the crude oil isseparated into the three fractions by a combination of distilling thecrude under vacuum, visbreaking vacuum distillation bottoms, andstripping gaseous by-products and additional sulfur-containingdistillate from the visbreaker effluent in a pitch stripper, forming thesulfur-containing residual fraction comprising pitch stripper bottoms.

3. A process in accordance with claim 1 in combination withcatalytically reforming a portion of sulfur-free distillate produced,and augmenting said high purity hydrogen with hydrogen produced in thecatalytic reforming.

4. A process which comprises (a) separating a sulfurcontaining crude oilinto a distillate fraction and a residual fraction; (b) partiallyoxidizing the entire residual fraction with molecular oxygen at anelevated pressure of at least 1500 p.s.i.g. and temperature conditionsfor synthesis gas generation in a hydrogen manufacturing zone, shiftconverting the CO in the gas produced to CO and removing H 5 and CO toprovide high purity hydrogen by contacting the shifted gas at theelevated pressure with an organic oxygen compound liquid absorbent at atemperature below F.; and (c) reacting the sulfurcontaining distillatefraction with a portion of said high purity hydrogen in contact with asulfactive hydrogenation catalyst at elevated temperature and pressurein a hydroconversion zone, converting sulfur in said distillate to H 8and forming hydrofined distillate and gaseous byproducts; and reactingat least a portion of said hydrofined distillate with the remainingportion of said high purity hydrogen in contact with an activehydrocracking catalyst at elevated temperature and pressure in ahydroconversion Zone, converting hydrofined distillate to hydrocrackedlower boiling distillate and forming gaseous byproducts.

5. A process in accordance with claim 4 wherein in the hydrogenmanufacturing Zone a portion of the CO produced is removed bycondensation from the shifted gas, and the remaining portion is removedby contacting with the organic oxygen compound.

References Cited UNITED STATES PATENTS 3,008,895 11/1961 Hansford et al.20868 3,012,962 12/1961 Dygert 208154- 3,120,993 2/1964 Thormann et al232 3,380,910 4/1968 Grifiiths 20858 2,911,352 11/1959 Goretta et al.20868 3,240,694 11/ 1963 Mason et al. 20859 3,254,020 5/1966 Prayer etal. 20889 DELBERT E. GANTZ, Primary Examiner T. H. YOUNG, AssistantExaminer U.S. Cl. X11, 2()8--58, 92

UNITED STATES PATENT OFFICE {56g CERTIFICATE OF CORRECTIQN Patent3.5371977 Dated November 3 1970 Inventor-( CALVIN 8. SMITH, JR.

It is certified that error appears in the above-identified patent andthat said Letters Patent are hereby corrected as shown below:

Col. 2, line 68, "Table II" should read SIGNED m smm Us 1971 B Am Mum!"mun. 5mm, 3 mg Offimr domhlioner of Patm'rt:-

